Aromatic Transalkylation Using A Modifed LZ-210 Zeolite

ABSTRACT

A process for converting polyalkylaromatics to monoalkylaromatics, particularly cumene, in the presence of a modified LZ-210 zeolite catalyst is disclosed. The process attains greater selectivity, reduced formation of undesired byproducts, and increased activity.

CROSS-REFERENCE TO RELATED APPLICATION

This application is a Continuation-In-Part of copending application Ser.No. 11/622,925 filed Jan. 12, 2007, and which is incorporated byreference in its entirety.

TECHNICAL FIELD

The process and catalyst disclosed herein relate to the production ofmonoalkylaromatics, in particular cumene, from polyalkylaromatics, inparticular polyisopropylbenzenes (PIPBs) including, but not necessarilylimited to, triisopropylbenzene (TIPB) and diisopropylbenzene (DIPB).The process relates to the use of a modified LZ-210 zeolite derived froma Y zeolite as a catalyst in the transalkylation of suchpolyalkylaromatics.

BACKGROUND

The following description will make specific reference to the use of thecatalyst disclosed herein in the transalkylation of PIPBs with benzeneto afford cumene, but it is to be recognized that this is done solelyfor the purpose of clarity and simplicity of exposition. Frequentreference will be made herein to the broader scope of this applicationfor emphasis.

Cumene is a major article of commerce, with one of its principal usesbeing a source of phenol and acetone via its air oxidation and asubsequent acid-catalyzed decomposition of the intermediatehydroperoxide.

Because of the importance of both phenol and acetone as commoditychemicals, there has been much emphasis on the preparation of cumene andthe literature is replete with processes for its manufacture. The mostcommon and perhaps the most direct method of preparing cumene is thealkylation of benzene with propylene, especially using an acid catalyst.

Another common method of preparing cumene is the transalkylation ofbenzene with PIPB, particularly di-isopropylbenzene (DIPB) andtri-isopropylbenzene (TIPB), especially using an acid catalyst. Anycommercially feasible transalkylation process must satisfy therequirements of a high conversion of polyalkylated aromatics and a highselectivity to monoalkylated products.

The predominant orientation of the reaction of benzene with PIPBresulting in cumene corresponds to Markownikoff addition of the propylgroup. However, a small but very significant amount of the reactionoccurs via anti-Markownikoff addition to afford n-propylbenzene (NPB).The significance of NPB formation is that it interferes with theoxidation of cumene to phenol and acetone, and consequently cumene usedfor oxidation must be quite pure with respect to NPB content.

Because cumene and NPB are difficult to separate by conventional means(e.g. distillation), the production of cumene via the transalkylation ofbenzene with PIPB must be carried out with a minimal amount of NPBproduction. One important factor to take into consideration is that theuse of an acid catalyst for the transalkylation results in increased NPBformation with increasing temperature. Thus, to minimize NPB formation,the transalkylation should be carried out at as low a temperature aspossible.

Since DIPB and TIPB are not only the common feeds for thetransalkylation of benzene with PIPBs but also the common byproducts ofthe alkylation of benzene with propylene when forming cumene,transalkylation is commonly practiced in combination with alkylation tominimize the production of less valuable byproducts and to produceadditional cumene. In such a combination process, the cumene produced byboth alkylation and transalkylation is typically recovered in a singleproduct stream. Since NPB is also formed in alkylation and the amount ofNPB formation in alkylation increases with increasing temperature, theNPB production in both alkylation and transalkylation must be managedrelative to one another so that the cumene product stream is relativelyfree of NPB.

What is needed is an optimum transalkylation catalyst for, e.g., cumeneproduction, with sufficient activity to effect transalkylation atacceptable reaction rates at temperatures sufficiently low to avoidunacceptable NPB formation. Because Y zeolites show substantiallygreater activity than many other zeolites, they have received closescrutiny as a catalyst in as a catalyst in aromatic transalkylation.However, a problem exists in that Y zeolites effect transalkylation atunacceptably low rates at the low temperatures desired to minimize NPBformation.

Therefore, in order for a commercial process based on Y zeolites tobecome a reality, it is necessary to increase catalyst activity—i.e.,increase the rate of cumene production at a given, lower temperature.

BRIEF SUMMARY OF THE DISCLOSURE

Aromatic transalkylation catalysts and processes disclosed hereininclude a modified LZ-210 zeolite. The catalysts and processes mayprovide, for example, decreased NPB formation and increased activityrelative to other Y zeolites.

Accordingly, in an embodiment the invention is an aromatictransalkylation catalyst comprising from about 5 wt % to about 90 wt %of a modified LZ-210 zeolite on a volatile-free basis, and an aluminabinder; the modified LZ-210 zeolite having a bulk Si/Al₂ molar ratioranging from about 6.5 to about 27; wherein 63% to 95% of aluminum inthe modified LZ-210 zeolite is framework aluminum. In anotherembodiment, the invention is an aromatic transalkylation catalystcomprising from about 5 wt % to about 90 wt % of a modified LZ-210zeolite on a volatile-free basis, and an alumina binder; the modifiedLZ-210 zeolite having an absolute intensity as measured by X-raydiffraction of at least 50; wherein 63% to 95% of aluminum in themodified LZ-210 zeolite is framework aluminum.

In other embodiments, the invention is a process for transalkylatingaromatics, the process comprising passing a transalkylatable aromaticand an aromatic substrate to a reaction zone; contacting thetransalkylatable aromatic and the aromatic substrate with atransalkylation catalyst in the reaction zone at transalkylationconditions including a temperature from about 100° C. to about 390° C.,a pressure of between 1 and 130 atmospheres, and a water concentrationof less than 20 wt-ppm based on the combined weight of thetransalkylatable aromatic and the aromatic substrate passed to thereaction zone.

Other embodiments of the process disclosed herein are described in thedetailed description.

BRIEF DESCRIPTION OF THE DRAWINGS

FIG. 1 illustrates, graphically, DIPB conversion (y-axis, %) versustemperature (x-axis, ° C.) for catalysts prepared in accordance withExamples 4-6 of this disclosure against Comparative Example 1;

FIG. 2 illustrates, graphically, the ratio of NPB to cumene (y-axis,wt-ppm) in the product versus DIPB conversion (x-axis, %) for thecatalysts of Examples 4-6 of this disclosure and against ComparativeExample 1;

FIG. 3 illustrates, graphically, DIPB conversion (y-axis, %) versustemperature (x-axis, ° C.) for the catalyst of Example 4 beforeregeneration and after regeneration (Example 9) and against ComparativeExample 1; and

FIG. 4 illustrates, graphically, the ratio of NPB to cumene (y-axis,wt-ppm) in the product versus DIPB conversion (x-axis, %) for thecatalyst of Example 4 before regeneration and after regeneration(Example 9) and against Comparative Example 1.

DETAILED DESCRIPTION

The process disclosed herein uses a catalyst that comprises acrystalline zeolitic molecular sieve. The preferred molecular sieves foruse in the catalyst disclosed herein are modified Y zeolites. U.S. Pat.No. 3,130,007, which is hereby incorporated herein by reference in itsentirety, describes Y-type zeolites. The modified Y zeolites suitablefor use in preparing the catalyst disclosed herein are generally derivedfrom Y zeolites by treatment which results in a significant modificationof the Y zeolite framework structure and composition, usually anincrease in the bulk Si/Al₂ mole ratio to a value typically above 6.5and/or a reduction in the unit cell size. It will be understood,however, that, in converting a Y zeolite starting material to a modifiedY zeolite useful in the process disclosed herein, the resulting modifiedY zeolite may not have exactly the same X-ray powder diffraction patternfor Y zeolites as described in the '007 patent. The modified Y zeolitemay have an X-ray powder diffraction pattern similar to that of the '007patent but with the d-spacings shifted somewhat due, as those skilled inthe art will realize, to cation exchanges, calcinations, etc., which aregenerally necessary to convert the Y zeolite into a catalytically activeand stable form.

The modified Y zeolite useful in the process disclosed herein has a unitcell size of from about 24.34 to about 24.58 Å, preferably from about24.36 to about 24.55 Å. The modified Y zeolite has a bulk Si/Al₂ molarratio of from about 6.5 to about 27, more preferably from about 6.5 toabout 23.

In preparing the modified Y zeolite component of the catalysts used inthe process described herein, the starting material may be a Y zeolitein alkali metal (e.g., sodium) form such as described in the '007patent. The alkali metal form Y zeolite is ion-exchanged with ammoniumions, or ammonium ion precursors such as quaternary ammonium or othernitrogen-containing organic cations, to reduce the alkali metal contentto less than about 4 wt %, preferably less than about 3 wt %, morepreferably less than about 2.5 wt %, expressed as the alkali metal oxide(e.g., Na₂O) on a dry basis. As used herein, the weight of the zeoliteon a water-free, volatile free, or dry basis means the weight of thezeolite after maintaining the zeolite at a temperature of about 900° C.(1652° F.) for roughly 2 hours.

Optionally, the starting zeolite can also contain or at some stage ofthe modification procedure be ion-exchanged to contain rare earthcations to the degree that the rare earth content as RE₂O₃ constitutesfrom about 0.1 to about 12.5 wt % of the zeolite (anhydrous basis),preferably from about 8.5 to about 12 wt %. It will be understood bythose skilled in the art that the ion-exchange capacity of the zeolitefor introducing rare earth cations decreases during the course of thedisclosed treatment process. Accordingly, if rare earth cation exchangeis carried out, for example, as the final step of the preparativeprocess, it may not be possible to introduce even the preferred amountof rare earth cations. The framework Si/Al₂ ratio of the starting Yzeolite can be within the range of less than about three 3 to about 6,but is advantageously greater than about 4.8.

Y zeolites which may be used in the process disclosed herein may beprepared by dealuminating a Y zeolite having an overall silica toalumina mole ratio below about 5 and are described in detail in U.S.Pat. Nos. 4,503,023, 4,597,956, 4,735,928 and 5,275,720 which are herebyincorporated herein by reference. The '023 patent discloses anotherprocedure for dealuminating a Y zeolite involving contacting the Yzeolite with an aqueous solution of a fluorosilicate salt usingcontrolled proportions, temperatures, and pH conditions which avoidaluminum extraction without silicon substitution. The '023 patentdiscloses that the fluorosilicate salt is used as the aluminumextractant and also as the source of extraneous silicon which isinserted into the Y zeolite structure in place of the extractedaluminum. The salts have the general formula:

(A)_(2/b)SiF₆

wherein A is a metallic or nonmetallic cation other than H⁺ having thevalence “b.” Cations represented by “A” are alkylammonium, NH₄ ⁺, Mg⁺⁺,Li⁺, Na⁺, K⁺, Ba⁺⁺, Cd⁺⁺, Cu⁺⁺, H⁺, Ca⁺⁺, Cs⁺, Fe⁺⁺, Co⁺⁺, Pb⁺⁺, Mn⁺⁺,Rb⁺, Ag⁺, Sr⁺⁺, Ti⁺, and Zn⁺⁺.

A preferred member of this group of Y zeolites is known as LZ-210, azeolitic aluminosilicate molecular sieve described in the '023 patent.LZ-210 zeolites and the other zeolites of this group are convenientlyprepared from a Y zeolite starting material. In one embodiment, theLZ-210 zeolite has an overall silica to alumina mole ratio from about5.0 to about 11.0. The unit cell size ranges from about 24.38 to about24.50 angstrom, preferably from about 24.40 to about 24.44 angstrom. TheLZ-210 class of zeolites used in the process and composition disclosedherein have a composition expressed in terms of mole ratios of oxides asin the following formula:

(0.85-1.1)M_(2/n)O:Al₂O₃ :xSiO₂

wherein “M” is a cation having the valence “n” and “x” has a value from5.0 to 11.0.

In general, LZ-210 zeolites may be prepared by dealuminating Y-typezeolites using an aqueous solution of a fluorosilicate salt, preferablya solution of ammonium hexafluorosilicate. The dealumination can beaccomplished by placing a Y zeolite, normally but not necessarily anammonium exchanged Y zeolite, into an aqueous reaction medium such as anaqueous solution of ammonium acetate, and slowly adding an aqueoussolution of ammonium fluorosilicate. After the reaction is allowed toproceed, a zeolite having an increased overall silica to alumina moleratio is produced. The magnitude of the increase is dependent at leastin part on the amount of fluorosilicate solution contacted with thezeolite and on the reaction time allowed. Normally, a reaction time ofbetween about 10 and about 24 hours is sufficient for equilibrium to beachieved. The resulting solid product, which can be separated from theaqueous reaction medium by conventional filtration techniques, is a formof form of LZ-210 zeolite. In some cases this product may be subjectedto a steam calcination by methods well known in the art. For instance,the product may be contacted with water vapor at a partial pressure ofat least 1.4 kPa(a) (0.2 psi(a)) for a period of between about ¼ toabout 3 hours at a temperature between about 482° C. (˜900° F.) andabout 816° C. (˜1500° F.) in order to provide greater crystallinestability. In some cases the product of the steam calcination may besubjected to an ammonium-exchange by methods well known in the art. Forinstance, the product may be slurried with water after which an ammoniumsalt is added to the slurry. The resulting mixture is typically heatedfor a period of hours, filtered, and washed with water. Methods ofsteaming and ammonium-exchanging LZ-210 zeolite are described in U.S.Pat. Nos. 4,503,023, 4,735,928, and 5,275,720.

In one embodiment, the ammonium exchange is followed by the treatmentwith an aqueous solution of a fluorosilicate salt to increase Si/Al₂ratio, enhancing the hydrothermal stability and lowering the propensityto form extra-framework aluminum.

The final low pH, ammonium ion exchange of the LZ-210 zeolite, which ispreferred, can be carried out in the same manner as in the case of theinitial ammonium exchange of the Y zeolite (and/or LZ-210 zeolite asdiscussed above) except that the pH of the exchange medium is lowered tobelow about 4, preferably to below about 3, at least during some portionof the ion-exchange procedure. The lowering of the pH is readilyaccomplished by the addition of an appropriate mineral or organic acidto the ammonium ion solution. Nitric acid and sulfuric are especiallysuitable for this purpose. Preferably, acids which form insolublealuminum salts are avoided. In performing the low pH ammonium ionexchange, both the pH of the exchange medium, the quantity of exchangemedium relative to the zeolite and the time of contact of the zeolitewith the exchange medium are significant factors. It is found that solong as the exchange medium is at a pH below 4, sodium cations areexchanged for hydrogen cations in the zeolite and, in addition, at leastsome aluminum, predominately non-framework and some framework, isextracted. The efficiency of the process is improved, however, byacidifying the ion exchange medium using more acid than is required tolower the pH to just below 4. As will be evident from the data set forthbelow, the more acidic the exchange medium is, the greater the tendencyto extract framework as well as non-framework aluminum from the zeolite.The extraction procedure is carried out to a degree sufficient toproduce a zeolite product having a bulk Si/Al₂ molar ratio ranging fromabout 6.5 to about 35. about 6.5 to about 35. In other embodiments, thebulk Si/Al₂ molar ratio ranges from about 6.5 to about 30, morepreferably from about 6.5 to about 27, still more preferably from about6.5 to about 23, or even more preferably from about 6.5 to about 20.

Although the disclosed catalyst may contain a metal hydrogenationcatalytic component, such a component is not a requirement. Based on theweight of the catalyst, such a metal hydrogenation catalytic componentmay be present at a level of less than 0.2 wt % or less than 0.1 wt %calculated as the respective oxide of the metal component, or thecatalyst may be devoid of any metal hydrogenation catalytic component.If present, the metal hydrogenation catalytic component can exist withinthe final catalyst composite as a compound such as an oxide, sulfide,halide and the like, or in the elemental metallic state. As used herein,the term “metal hydrogenation catalytic component” is inclusive of thesevarious compound forms of the metals. The catalytically active metal canbe contained within the inner adsorption region, i.e., pore system, ofthe zeolite constituent, on the outer surface of the zeolite crystals orattached to or carried by a binder, diluent or other constituent, ifsuch is employed. The metal can be imparted to the overall compositionby any method which will result in the attainment of a highly dispersedstate. Among the suitable methods are impregnation, adsorption, cationexchange, and intensive mixing. The metal can be copper, silver, gold,titanium, chromium, molybdenum, tungsten, rhenium, manganese, zinc,vanadium, or any of the elements in IUPAC Groups 8-10 especiallyplatinum, palladium, rhodium, cobalt, and nickel. Mixtures of metals maybe employed.

The finished catalyst compositions can contain the usual binderconstituents in amounts which are in the range of from about 10 to about95 wt %, preferably from about 15 to 50 wt %. The binder is ordinarilyan inorganic oxide or mixtures thereof. Both amorphous and crystallinecan be employed. Examples of suitable binders are silica, alumina,silica-alumina, clays, zirconia, silica-zirconia and silica-boria.Alumina is a preferred binder material.

For cumene production, the finished catalyst, made of 80 wt % zeoliteand 20 wt % alumina binder on a volatile-free basis, preferably has one,and more preferably both, of the following characteristics: (1) anabsolute intensity of the modified LZ-210 zeolite as measured by X-raydiffraction (XRD) of preferably at least 50, more preferably at least60; and and (2) a framework aluminum of the modified LZ-210 zeolite ofpreferably at least 60%, more preferably at least 70%, of the aluminumof the modified LZ-210 zeolite. In an embodiment, 63% to 95% of aluminumin the modified LZ-210 zeolite is framework aluminum. In one example,the finished catalyst for cumene production has a product of theabsolute intensity of the modified LZ-210 zeolite as measured by XRD andthe % framework aluminum of the aluminum in the modified LZ-210 zeolitethat is greater than 4200. For ethylbenzene production, the finishedcatalyst preferably has one, and more preferably both, of the followingcharacteristics: (1) an absolute intensity of the modified LZ-210zeolite as measured by X-ray diffraction (XRD) of preferably at least65, more preferably at least 75; and (2) a framework aluminum of themodified LZ-210 zeolite of preferably at least 50%, more preferably atleast 60%, of the aluminum of the modified LZ-210 zeolite. In anembodiment, 63% to 95% of aluminum in the modified LZ-210 zeolite isframework aluminum. In one example, the finished catalyst forethylbenzene production has a product of the absolute intensity of themodified LZ-210 zeolite as measured by XRD and the % framework aluminumof the aluminum in the modified LZ-210 zeolite that is greater than4500. As illustrated in FIGS. 1-4 and the examples below, the disclosedcatalysts provide increase catalyst activity and, in the case of cumeneproduction, lower NPB formation. In the case of ethylbenzene productionfrom poly-ethylbenzenes (FIG. 5), while internal isomerization of ethylgroups is of little concern and even though an ethyl group is smallerthan a propyl group, the diffusion characteristics of the disclosedcatalysts appear to be important.

In one embodiment, the process disclosed herein uses a catalyst that issubstantially dry. The low pH, ammonium ion exchange is not necessarilyfollowed by a calcination step that drives off substantially all of thewater present. It has been found that the performance of the catalyst inthe process described herein is improved by removing water. In order tomaintain high activity and low NPB formation, it has been found that thewater content of the zeolite must be relatively low before it is used inthe transalkylation process.

Excess water may reduce the number of active sites and restrictdiffusion to them so they do not efficiently catalyze transalkylation.To address this problem, dehydration of the catalyst particles so theycontain the desired amount of water may be carried out, prior tostart-up, with a drying agent that may be introduced into thetransalkylation reaction zone, as the the temperature in the reactionzone may be slowly increased to before the aromatic substrate or thetransalkylatable aromatic is introduced. During this initial heat-upperiod, the water content of the zeolite is determined by theequilibrium between the zeolite, the catalyst, the drying agent, and theamount of water in the reaction zone, if any, at temperatures in thereaction zone. The zeolitic portion of the catalyst is highlyhydrophilic and the level of hydration is controlled by adjusting therate at which the drying agent passes over the catalyst and thetemperature during the dehydration step. The drying agent may be anyagent that removes water and does not have a deleterious effect on thecatalyst, such as molecular nitrogen, air, or benzene. The temperatureduring the dehydration step is maintained between about 25 and about500° C. (˜77 to ˜932° F.). The water content of the catalyst iscalculated by measuring weight loss on ignition (LOI), which is normallydetermined by calculating the weight loss after heating for about 2hours at about 900° C. (˜1652° F.), and then subtracting the amount ofweight loss due to ammonium ion decomposition into ammonia. Since acatalyst containing water in excess of the desired amount, i.e., greaterthan the equilibrium amount of water the catalyst will contain at anytime during process start-up, will lose water once equilibrium isestablished during start-up, it is not necessary, though it may bedesirable, for the dehydration step to be carried out to give thecatalyst an amount of water that is equal to or less than theequilibrium amount.

Some desired properties of the catalyst, such as crush strength andammonium ion concentration, are achieved by controlling the time andtemperature conditions at which the extruded catalyst particles arecalcined. In some cases, calcination at higher temperatures will leavethe required amount of water in the catalyst and thereby make itunnecessary to carry out a separate dehydration step. Thus,“dehydrating” and “dehydration” as used herein not only mean a separatestep in which water is removed to the catalyst after calcination butalso encompass a calcination step carried out under conditions such thatthe desired amount of water remains on the catalyst particles.

The dehydration procedure described above is part of the actual processof making the disclosed catalyst at the manufacturing plant. It will beunderstood, however, that procedures other than that described above canbe used to dehydrate the catalyst either in the manufacturing plant atthe time the catalyst is made or at some other time at the manufacturingplant or elsewhere. For example, the extruded catalyst particles can bedehydrated in-situ in dehydrated in-situ in the transalkylation reactorby passing a water-deficient containing gas, such as dry molecularnitrogen or air, or a dry reactant, such as dry aromatic substrate(e.g., benzene) or dry transalkylatable aromatic (e.g., DIPB or TIPB),over the catalyst at relatively high temperatures until the catalystcontains the desired amount of water. In an in-situ dehydration step,the water-deficient gas or reactant typically contains less than about30 wt-ppm water, and the contacting is done at a temperature betweenabout 25° C. (˜77° F.) to about 500° C. (˜932° F.). In one example, thecatalyst is contacted with flowing dry nitrogen in the gas phase atabout 250° C. (˜482° F.). The catalyst is contacted with flowing drybenzene in the liquid phase at, for example, about 120° C. (˜248° F.) toabout 260° C. (˜500° F.), about 160° C. (˜320° F.) to about 210° C.(˜410° F.), about 180° C. (˜356° F.) to about 200° C. (˜392° F.), orabout 120° C. (˜248° F.) to about 180° C. (˜356° F.). Also, the catalystparticles can be stored at the manufacturing plant or elsewhere so thatthey are in contact with a surrounding gas until the desired amount ofwater has been desorbed.

Typically, the LOI of the catalyst that is loaded into thetransalkylation reactor is in the range of from about 2 to about 4 wt %.After loading in the reactor, and preferably prior to using the catalystto promote transalkylation reactions, the catalyst may be subjected tothe dehydration step to decrease the water content of the catalyst. Thenitrogen content of the catalyst is also preferably minimized.

The disclosed catalyst is useful in the transalkylation oftransalkylatable aromatics. The transalkylation process disclosed hereinpreferably accepts as feed a transalkylatable hydrocarbon (i.e. apolyalkylaromatic) in conjunction with an aromatic substrate. Thetransalkylatable hydrocarbons useful in the transalkylation process arecomprised of aromatic compounds which are characterized as constitutingan aromatic substrate based molecule with one or more alkylating agentcompounds taking the place of one or more hydrogen atoms around thearomatic substrate ring structure. For example, the transalkylatablearomatic may include one or more polyisopropylbenzene such asdiisopropylbenzene (DIPB) and triisopropylbenzene (TIPB) in theproduction of cumene. In an embodiment to produce ethylbenzene, thetransalkylatable aromatic may include one or more polyethylbenzene suchas diethylbenzene, triethylbenzene, and tetraethylbenzene.

The alkylating agent compounds which may be selected from a group ofdiverse materials including monoolefins, diolefins, polyolefins,acetylenic hydrocarbons, and also alkylhalides, alcohols, ethers esters,the later including the alkylsulfates, alkylphosphates and variousesters of carboxylic acids. The preferred olefin-acting compounds areolefinic hydrocarbons which comprise monoolefins containing one doublebond per molecule. Monoolefins which may be utilized as olefin-actingcompounds in the disclosed process are either normally gaseous ornormally liquid and include ethylene, propylene, 1-butene, 2-butene,isobutylene, and the high molecular weight normally liquid olefins suchas the various pentenes, hexenes, heptenes, octenes, and mixturesthereof, and still higher molecular weight liquid olefins, the latterincluding various olefin oligomers having from about 9 to about 18carbon atoms per molecule including propylene trimer, propylenetetramer, propylene pentamer, etc. C₉ to C₁₈ normal olefins may be usedas may cycloolefins such as cyclopentene, methylcyclopentene,cyclohexene, methylcyclohexene, etc. may also be utilized, although notnecessarily with equivalent results. It is that the monoolefin containsat least 2 and not more than 14 carbon atoms. More specifically, it ispreferred that the monoolefin is propylene. The alkylating agentcompounds are preferably C₂-C₁₄ aliphatic hydrocarbons, and morepreferably propylene.

The aromatic substrate useful as a portion of the feed to thetransalkylation process may be selected from a group of aromaticcompounds which include individually and in admixture with benzene andmonocyclic alkylsubstituted benzene having the structure:

where R is a hydrocarbon containing 1 to 14 carbon atoms, and n is aninteger from 1 to 5. In other words, the aromatic substrate portion ofthe feedstock may be benzene, benzene containing from 1 to 5 alkyl groupsubstituents, and mixtures thereof. Non-limiting examples of suchfeedstock compounds include benzene, toluene, xylene, ethylbenzene,mesitylene (1,3,5-trimethylbenzene), cumene, n-propylbenzene,butylbenzene, dodecylbenzene, tetradecylbenzene, and mixtures thereof.It is specifically preferred that the aromatic substrate is benzene.

The disclosed transalkylation process may have a number of purposes. Inone, the catalyst of the transalkylation reaction zone is utilized toremove the alkylating agent compounds in excess of one from the ringstructure of polyalkylated aromatic compounds and to transfer thealkylating agent compound to an aromatic substrate molecule that has notbeen previously alkylated, thus increasing the amount of the desiredaromatic compounds produced by the process. In a related purpose, thereaction performed in the transalkylation reaction zone involves theremoval of all alkylating agent components from a substituted aromaticcompound and in doing so, converting the aromatic substrate intobenzene.

The feed mixture has a concentration of water and oxygen-containingcompounds in the combined feed of preferably less than about 20 wt-ppm,more preferably less than about 10 wt-ppm, and yet more preferably lessthan about 2 wt-ppm based on the weight of the transalkylatable aromaticand an aromatic substrate passed to the reaction zone. The method bywhich such low concentrations in the feed mixture are attained is notcritical to the process disclosed herein. Usually, one stream containingthe transalkylatable aromatic and another stream containing the aromaticsubstrate are provided, with each stream having a concentration of waterand oxygen-containing compounds precursors such that the feed mixtureformed by combining the individual streams has the desiredconcentration. Water and oxygen-containing compounds can be removed fromeither the individual streams or the feed mixture by conventionalmethods, such as drying, adsorption, or stripping. Oxygen-containingcompounds may be any alcohol, aldehyde, epoxide, ketone, phenol or etherthat has a molecular weight or boiling point within the range ofmolecular weights or boiling points of the hydrocarbons in the feedmixture.

To transalkylate polyalkylaromatics with an aromatic substrate, a feedmixture containing an aromatic substrate and polyalkylated aromaticcompounds in mole ratios ranging from about 1:1 to about 50:1 andpreferably from about 1:1 to about 10:1 are continuously orintermittently introduced into a transalkylation reaction zonecontaining the disclosed catalyst at transalkylation conditionsincluding a temperature from about 100 to about 390° C. (˜212 to ˜734°F.), and especially from about 110 to about 275° C. (˜230 to ˜527° F.).Pressures which are ˜527° F.). Pressures which are suitable for useherein preferably are above 1 atmosphere (101.3 kPa(a)) but should notbe in excess of about 130 atmospheres (13169 kPa(a)). An especiallydesirable pressure range is from about 10 to about 40 atmospheres (˜1013to ˜4052 kPa(a)). A weight hourly space velocity (WHSV) of from about0.1 to about 50 hr⁻¹, and especially from about 0.5 to about 5 hr⁻¹,based upon the polyalkylaromatic feed rate and the total weight of thecatalyst on a dry basis, is desirable. While the process disclosedherein may be performed in the vapor phase, it should be noted that thetemperature and pressure combination utilized in the transalkylationreaction zone is preferred to be such that the transalkylation reactionstake place in essentially the liquid phase. In a liquid phasetransalkylation process for producing monoalkylaromatics, the catalystis continuously washed with reactants, thus preventing buildup of cokeprecursors on the catalyst. This results in reduced amounts of carbonforming on said catalyst in which case catalyst cycle life is extendedas compared to a gas phase transalkylation process in which cokeformation and catalyst deactivation is a major problem. Additionally,the selectivity to monoalkylaromatic production, especially cumeneproduction, is higher in the catalytic liquid phase transalkylationreaction herein as compared to catalytic gas phase transalkylationreaction.

Transalkylation conditions for the process disclosed herein include amolar ratio of aromatic ring groups per alkyl group of generally fromabout 1:1 to about 25:1. The molar ratio may be less than 1:1, and it isbelieved that the molar ratio may be 0.75:1 or lower. Preferably, themolar ratio of aromatic ring groups per alkyl propyl group (or perpropyl group, in cumene production) is below 6:1.

At transalkylation conditions, the catalyst particles typically containwater in an amount preferably below about 4 wt %, more preferably belowabout 3 wt %, and yet more preferably below about 2 wt %, as measured byKarl Fischer titration, and nitrogen in an amount preferably below about0.05 wt %, as measured by micro (CHN) (carbon-hydrogen-nitrogen)analysis.

All references herein to the groups of elements of the periodic tableare to the IUPAC “New Notation” on the Periodic Table of the Elements inthe inside front cover of the book entitled CRC Handbook of Chemistryand Physics, ISBN 0-8493-0480-6, CRC Press, Boca Raton, Fla., U.S.A.,80^(th) Edition, 1999-2000.

As used herein, the molar ratio of aromatic ring groups per alkyl groupis defined as follows. The numerator of this ratio is the number ofmoles of aromatic ring groups passing through the reaction zone during aspecified period of time. The number of moles of aromatic ring groups isthe sum of all aromatic ring groups, regardless of the compound in whichthe aromatic ring group happens to be. For example, in cumene productionone mole of benzene, one mole of cumene, one mole of DIPB, and one moleof TIPB each contribute one mole of aromatic ring group to the sum ofaromatic ring groups. In ethylbenzene (EB) production, one mole ofbenzene, one mole of EB, and one mole of di-ethylbenzene (DEB) eachcontribute one mole of aromatic ring group to the sum of aromatic ringgroups. The denominator of this ratio is the number of moles of alkylgroups that have the same number of carbon atoms as that of the alkylgroup on the desired monoalkylated aromatic and which pass through thereaction zone during the same specified period of time. The number ofmoles of alkyl groups is the sum of all alkyl and alkenyl groups withthe same number of carbon atoms as that of the alkyl group on thedesired monoalkylated aromatic, regardless of the compound in which thealkyl or alkenyl group happens to be, except that paraffins are notincluded. Thus, the number of moles of propyl groups is the sum of alliso-propyl, n-propyl, and propenyl groups, regardless of the compound inwhich the iso-propyl, n-propyl, or propenyl group happens to be, exceptthat paraffins, such as propane, n-butane, isobutane, pentanes, andhigher paraffins are excluded from the computation of the number ofmoles of propyl groups. For example, one mole of propylene, one mole ofcumene, and one mole of NPB each contribute one mole of propyl group tothe sum of propyl groups, whereas one mole of DIPB contributes two molesof propyl groups and one mole of tri-propylbenzene contributes threemoles of propyl groups regardless of the distribution of the threegroups between iso-propyl and n-propyl groups. One mole of ethylene andone mole of EB each contribute one mole of ethyl groups to the sum ofethyl groups, whereas one mole of DEB contributes two moles of ethylgroups and one mole of tri-ethylbenzene contributes three moles of ethylgroups. Ethane contributes no moles of ethyl groups.

As used herein, WHSV means weight hourly space velocity, which isdefined as the weight flow rate per hour divided by the catalyst weight,where the weight flow rate and the catalyst weight are in the sameweight units.

As used herein, DIPB conversion is defined as the difference between themoles of DIPB in the feed and the moles of DIPB in the product, dividedby the moles of DIPB in the feed, multiplied by 100.

All references herein to surface area are calculated using nitrogenpartial pressure p/po data points ranging from about 0.03 to about 0.30using the BET (Brunauer-Emmett-Teller) model method using nitrogenadsorption technique as described in ASTM D4365-95, Standard Test Methodfor Determining Micropore Volume and Zeolite Area of a Catalyst, and inthe article by S. Brunauer et al., J. Am. Chem. Soc., 60(2), 309-319(1938).

As referred to herein, the absolute intensity by X-ray powderdiffraction (XRD) of a Y zeolite material was measured by computing thenormalized sum of the intensities of a few selected XRD peaks of the Yzeolite material and dividing that sum by the normalized sum of theintensities of a few XRD peaks of the alpha-alumina NBS 674a intensitystandard, which is the primary standard and which is certified by theNational Institute of Standards and Technology (NIST), an agency of theU.S. Department of Commerce. The Y zeolite's absolute intensity is thequotient of the sums multiplied by 100:

${{Absolute}\mspace{14mu} {Intensity}} = \frac{\begin{pmatrix}{{Normalized}\mspace{14mu} {Intensity}\mspace{14mu} {of}} \\{Y\mspace{14mu} {Zeolite}\mspace{14mu} {Material}\mspace{14mu} {Peaks}}\end{pmatrix} \times 100}{\begin{pmatrix}{{{Normalized}\mspace{14mu} {Intensity}\mspace{14mu} {of}\mspace{14mu} {Alpha}} -} \\{{Alumina}\mspace{14mu} {Standard}\mspace{14mu} {Peaks}}\end{pmatrix}}$

The scan parameters of the Y zeolite material and the alpha-aluminastandard are shown in Table 1.

TABLE 1 Material Y zeolite Alpha-alumina standard 2T Ranges 4-5624.6-26.6, 34.2-36.2, 42.4-44.4 Step Time 1 sec/step or more 1 sec/stepdepending on zeolite content Step Width 0.02 0.01 Peaks (511,333),(440), (533), (012), (104), (113) (642), (751,555) + (660,822), (664)

For purposes of this disclosure, the absolute intensity of a Y zeolitethat is mixed with a nonzeolitic binder to give a mixture of Z parts byweight of the Y zeolite and (100−Z) parts by weight of the nonzeoliticbinder on a dry basis can be computed from the absolute intensity of themixture, using the formula, A=C·(100/Z), where A is the absoluteintensity of of the Y zeolite and C is the absolute intensity of themixture. For example, where the Y zeolite is mixed with HNO₃-peptizedPural SB alumina to give a mixture of 80 parts by weight of zeolite and20 parts by weight Al₂O₃ binder on a dry basis, and the measuredabsolute intensity of the mixture is 60, the absolute intensity of the Yzeolite is computed to be (60)·(100/80) or 75.

As used herein, the unit cell size, which is sometimes referred to asthe lattice parameter, means the unit cell size calculated using amethod which used profile fitting to find the XRD peak positions of the(642), (822), (555), (840) and (664) peaks of faujasite and the silicon(111) peak to make the correction.

As used herein, the bulk Si/Al₂ mole ratio of a zeolite is the silica toalumina (SiO₂ to Al₂O₃) mole ratio as determined on the basis of thetotal or overall amount of aluminum and silicon (framework andnon-framework) present in the zeolite, and is sometimes referred toherein as the overall silica to alumina (SiO₂ to Al₂O₃) mole ratio. Thebulk Si/Al₂ mole ratio is obtained by conventional chemical analysiswhich includes all forms of aluminum and silicon normally present.

As used herein, the fraction of the aluminum of a zeolite that isframework aluminum is calculated based on bulk composition and theKerr-Dempsey equation for framework aluminum from the article by G. T.Kerr, A. W. Chester, and D. H. Olson, Acta. Phys. Chem., 1978, 24, 169,and the article by G. T. Kerr, Zeolites, 1989, 9, 350.

As used herein, dry basis means based on the weight after drying inflowing air at a temperature of about 900° C. (˜1652° F.) for about 1hr.

The following examples are presented for purposes of illustration onlyand are not intended to limit the scope of this disclosure.

EXAMPLE 1 COMPARATIVE

A sample of Y-74 zeolite was slurried in a 15 wt % NH₄NO₃ aqueoussolution and the solution temperature was brought up to 75° C. (167°F.). Y-74 zeolite is a stabilized sodium Y zeolite with a bulk Si/Al₂ratio of approximately 5.2, a unit cell size of approximately 24.53, anda sodium content of approximately 2.7 wt % calculated as Na₂O on a drybasis. Y-74 74 zeolite is prepared from a sodium Y zeolite with a bulkSi/Al₂ ratio of approximately 4.9, a unit cell size of approximately24.67, and a sodium content of approximately 9.4 wt % calculated as Na₂Oon a dry basis that is ammonium exchanged to remove approximately 75% ofthe Na and then steam de-aluminated at approximately 600° C. (1112° F.)by generally following steps (1) and (2) of the procedure described incol. 4, line 47 to col. 5, line 2 of U.S. Pat. No. 5,324,877. Y-74zeolite is produced and was obtained from UOP LLC, Des Plaines, Ill.USA. After 1 hour of contact at 75° C. (167° F.), the slurry wasfiltered and the filter cake was washed with an excessive amount of warmde-ionized water. These NH₄ ⁺ ion exchange, filtering, and water washsteps were repeated two more times, and the resulting filter cake had abulk Si/Al₂ ratio of 5.2, a sodium content of 0.13 wt % calculated asNa₂O on a dry basis, a unit cell size of the 24.572 Å and an absoluteintensity of 96 as determined X-ray diffraction. The resulting filtercake was dried to an appropriate moisture level, mixed withHNO₃-peptized Pural SB alumina to give a mixture of 80 parts by weightof zeolite and 20 parts by weight Al₂O₃ binder on a dry basis, and thenextruded into 1.59 mm ( 1/16 in) diameter cylindrical extrudate. Theextrudate was dried and calcined at approximately 600° C. (1112° F.) forone hour in flowing air. This catalyst was representative of theexisting art. This catalyst had a unit cell size of 24.494 Å, an XRDabsolute intensity of 61.1, and 57.2% framework aluminum as a percentageof the aluminum in the modified Y zeolite.

EXAMPLE 2

As synthesized Y-54 zeolite was ammonium exchanged and then treated withammonium fluorosilicate according to the procedure described in U.S.Pat. No. 4,503,023. Y-54 zeolite is a sodium Y zeolite with a bulkSi/Al₂ ratio of approximately 4.9, a unit cell size of 24.67, and asodium content of 9.4 wt % calculated as Na₂O on a dry basis. Y-54zeolite is produced and was obtained from UOP LLC, Des Plaines, Ill.USA. The resulting Y zeolite, which had a bulk Si/Al₂ molar ratio ofabout 6.5, was steamed at about 600° C. (1112° F.) with 100% steam for 1hour, and then ammonium exchanged. The resulting filter cake was driedto an appropriate moisture level, mixed with HNO₃-peptized Pural SBalumina to give a mixture of 80 parts by weight of zeolite and 20 partsby weight Al₂O₃ binder on a dry basis, and then extruded into 1.59 mm (1/16 in) diameter cylindrical extrudate. The extrudate was dried andcalcined at approximately 600° C. (1112° F.) for one hour in flowingair. The hour in flowing air. The resulting catalyst had a unit cellsize of 24.426 Å, an absolute XRD intensity of 81.6, and 63.2% frameworkaluminum as a percentage of the aluminum in the modified Y zeolite.

EXAMPLE 3

As synthesized Y-54 zeolite was ammonium exchanged and then treated withammonium fluorosilicate according to the procedure described in U.S.Pat. No. 4,503,023. The resulting Y zeolite, which had a bulk Si/Al₂molar ratio of about 9.0 and was referred to as LZ-210(9), was steamedat about 600° C. (1112° F.) with 100% steam for 1 hour. A slurry made upof 228 g of the steamed LZ-210(9) and 672 g of H₂O was first prepared. ANH₄NO₃ solution made up of 212 g of H₂O and 667 g of 50 wt % (NH₄)NO₃was then added to the steamed LZ-210(9) slurry. The resulting mixturewas then raised to 85° C. (185° F.) and then mixed for 15 minutes. Tothis mixture, 5.7 g of 66 wt % HNO₃ were added, and the resultingmixture was maintained at 85° C. (185° F.) with continuous agitation for60 minutes. At the end of acid extraction, the mixture was filtered andthe cake was washed with 1000 ml of H₂O, and then dried at 100° C. (212°F.) overnight. In the second part, 200 g of dry cake was added to asolution made up of 667 g of 50 wt % (NH₄)NO₃ and 650 g of H₂O, to which20 g of 66 wt % HNO₃ was added. The resulting slurry was mixed for 60minutes. Thereafter, the mixture was filtered, washed with 1000 ml ofH₂O and the filter cake was oven dried at 100° C. (212° F.) overnight.The resulting zeolite had a 10.82 bulk Si/Al₂ ratio and 0.026 wt % Na₂O.The zeolite powder was mixed with HNO₃-peptized Pural SB alumina to givea mixture of 80 parts by weight of zeolite and 20 parts by weight Al₂O₃binder on a dry basis, moisture adjusted to give proper dough textureand then extruded into 1.59 mm ( 1/16 in) diameter cylindricalextrudate. The extrudate was dried and calcined at approximately 600° C.(1112° F.) for one hour in flowing air. The resulting catalyst had aunit cell size of 24.430 Å, an absolute XRD intensity of 78.4, 77.8%framework aluminum and a BET surface area of 661 m²/g.

EXAMPLE 4

As synthesized Y-54 zeolite was ammonium exchanged and then treated withammonium fluorosilicate according to the procedure described in U.S.Pat. No. 4,503,023. The resulting Y zeolite, which had a bulk Si/Al₂molar ratio of about 9.0 and was referred to as LZ-210(9), was steamedat about 600° C. (1112° F.) with 100% steam for 1 hour. An amount of 256g of the steamed LZ-210(9) was added to 1140 g of 22 wt % NH₄NO₃. To thezeolite slurry, 368 g of 17 wt % HNO₃ was slowly added over a period of30 minutes. The slurry was then heated up to 80° C. (176° F.) and heldat 80° C. (176° F.) for 90 minutes. At the end of acid extraction, theslurry was quenched with 1246 g of H₂O, filtered, washed with 1140 g ofa 22 wt % NH₄NO₃, washed with 1000 ml of H₂O and oven dried at 100° C.(212° F.) overnight. The resulting zeolite had a bulk 14.38 Si/Al₂ ratioand 0.047 wt % Na₂O. The resulting zeolite powder was mixed withHNO₃-peptized Pural SB alumina to give a mixture of 80 parts by weightof zeolite and 20 parts by weight Al₂O₃ binder on a dry basis, moistureadjusted to give proper dough texture and then extruded into 1.59 mm (1/16 in) diameter cylindrical extrudate. The extrudate was dried andcalcined at approximately 600° C. (1112° F.) for one hour in flowingair. The resulting catalyst had a unit cell size of 24.393 Å, anabsolute XRD intensity of 79.6, 81.8% framework aluminum, and a BETsurface area of 749 m²/g.

EXAMPLE 5

As synthesized Y-54 zeolite was ammonium exchanged and then treated withammonium fluorosilicate according to the procedure described in U.S.Pat. No. 4,503,023. The resulting Y zeolite, which had a bulk Si/Al₂molar ratio of about 12 and was referred to as LZ-210(12), was steamedat about 600° C. (1112° F.) with 100% steam for 1 hour. A slurry made upof 231 g of the steamed LZ-210(12) and 668 g of H₂O was first prepared.A NH₄NO₃ solution made up of 212 g of H₂O and 667 g of 50 wt % (NH₄)NO₃was then added to the steamed LZ-210(12) slurry. The resulting mixturewas then raised to 85° C. (185° F.) and then mixed for 15 minutes. Tothis mixture, 33.4 g of 66 wt % HNO₃ were added, and the resultingmixture was maintained at 85° C. (185° F.) with continuous agitation for60 minutes. At the end of acid extraction, the mixture was filtered andthe cake was washed with 1000 ml of H₂O, and of H₂O, and then dried at100° C. (212° F.) overnight. In the second part, 200 g of dry cake wasadded to a solution made up of 667 g of 50% (NH₄)NO₃ and 650 g of H₂O,to which 10 g of 66 wt % HNO₃ were added. The resulting slurry was mixedfor 60 minutes. Thereafter, the mixture was filtered, washed with 1000ml of H₂O and the filter cake was oven dried at 100° C. (212° F.)overnight. The resulting zeolite had a 17.24 bulk Si/Al₂ ratio and 0.01wt % Na₂O. The resulting zeolite powder was mixed with HNO₃-peptizedPural SB alumina to give a mixture of 80 parts by weight of zeolite and20 parts by weight Al₂O₃ binder on a dry basis, moisture adjusted togive proper dough texture and then extruded into 1.59 mm ( 1/16 in)diameter cylindrical extrudate. The extrudate was dried and calcined atapproximately 600° C. (1112° F.) for one hour in flowing air. Theresulting catalyst had a unit cell size of 24.391 Å, an absolute XRDintensity of 81.2, 94.9% framework aluminum and a BET surface area of677 m²/g.

EXAMPLE 6

An amount of 250 g of the LZ-210(12) from Example 5 (before steaming)was added to a NH₄NO₃ solution made up of 500 g of 50% NH₄NO₃ and 625 gof H₂O. The slurry was heated up to 95° C. (203° F.) and hold attemperature for 2 hours. The slurry was then filtered and water washed.The cake was then NH₄NO₃ exchanged and water washed a second timefollowing the same procedure. The filter cake was oven dried at 100° C.(212° F.) overnight. The resulting zeolite had a 12.62 bulk Si/Al₂ ratioand 0.05 wt % Na₂O. The dried zeolite was mixed with HNO₃-peptized PuralSB alumina to give a mixture of 80 parts by weight of zeolite and 20parts by weight Al₂O₃ binder on a dry basis, moisture adjusted to giveappropriate dough texture and then extruded into 1.59 mm ( 1/16 in)diameter cylindrical extrudate. The extrudate was dried and calcined atapproximately 600° C. (1112° F.) for one hour in flowing air. Theresulting catalyst had a unit cell size of 24.431 Å, an absolute XRDintensity of 77.3, 89.2% framework aluminum and a BET surface area of660 m²/g.

Table 2 summarizes the properties of the catalysts prepared in Examples1-6.

TABLE 2 Example 1 2 3 4 5 6 8 Type of Example Comparative ExampleExample Example Example Example Example Figures w/Run Data 1-4 None None1-4 1-2 1-2 1-2 Y zeolite bulk Si/Al₂ 5.20 8.61 10.82 14.38 17.24 12.6212.62 ratio, molar Y zeolite unit cell 24.494 24.426 24.430 24.39324.391 24.431 24.439 size, Å Catalyst XRD 61.1 81.6 78.4 79.6 81.2 77.372.5 absolute intensity Y zeolite XRD 76.4 102 98 99.5 101.5 96.6 90.6absolute intensity Y zeolite framework 57.2 63.2 77.8 81.8 94.9 89.292.6 aluminum, atomic % of total aluminum Catalyst BET surface — — 661749 677 660 660 area, m²/g

EXAMPLE 7

The catalysts prepared in the Examples 1 and 4-6 were tested fortransalkylation performance using a feed containing benzene andpolyalkylated benzenes. The feed was prepared by blending polyalkylatedbenzenes obtained from a commercial transalkylation unit with benzene.The feed composition as measured by gas chromatography is summarized inTable 3. The test was done in a fixed bed reactor in a once-through modeunder conditions of 3447 kPa(g) (500 psi(g)) reactor pressure, a molarratio of aromatic ring groups per propyl group of 2.3, and a 0.8 hr⁻¹DIPB WHSV over a range of reaction temperatures. The reactor was allowedto achieve essentially steady-state conditions at each reactiontemperature, and the product was sampled for analysis. Essentially nocatalyst deactivation occurred during the test. Prior to introducing thefeed, each catalyst was subjected to a drying procedure by contactingwith a flowing nitrogen stream containing less than 10 wt-ppm water at250° C. (482° F.) for 6 hours.

TABLE 3 Component Concentration, wt % Benzene 63.832 Nonaromatics 0.038Toluene 0.002 Ethylbenzene 0.000 Cumene 0.880 NPB 0.002 Butylbenzene0.071 Pentylbenzene 0.021 m-DIPB 20.776 o-DIPB 0.520 p-DIPB 13.472Hexylbenzene 0.308 1,3,5-TIPB 0.029 1,2,4-TIPB 0.012Tetra-isopropylbenzene 0.003 Nonylbenzene 0.004 Unknowns 0.030 Total100.000

FIGS. 1 and 2 show the test results for the catalysts prepared inExamples 1 and 4-6. In FIG. 1, the catalysts prepared in Examples 4-6show higher activities (i.e., higher DIPB conversion at a giventemperature) as compared to the curve 101 for Example 1. In FIG. 2, thecatalysts prepared in Examples 4-6 also exhibit better product purities(i.e., lower NPB/cumene at a given DIPB conversion) than the curve 201for the catalyst prepared in Example 1. Referring to FIGS. 1 and 2, thedata for Example 6 indicates that the steaming and acid extraction stepsare not required in the catalyst preparation, since good performance canbe achieved even when both are omitted. Still referring to FIGS. 1 and2, the data for Example 4 indicates that superior activity andcomparable product purity can be achieved using a single-steppost-steaming acid extraction, instead of the two-step acid extractionof Example 5, despite the acid extraction conditions being more severe.

EXAMPLE 8

A sample of the catalyst prepared in Example 6 was tested in the mannerdescribed in Example 7, as described previously. After testing, thespent catalyst was placed in a ceramic dish, which was placed in amuffle furnace. While flowing air was passed through the muffle furnace,the furnace temperature was raised from 70° C. (158° F.) to 550° C.(1022° F.) at a rate of 1° C. (1.8° F.) per minute, held at 550° C.(1022° F.) for 6 hours, and then cooled to 110° C. (230° F.). Theregenerated catalyst had a unit cell size of 24.439 Å, an absolute XRDintensity of 72.5, 92.6% framework aluminum and a BET surface area of660 m²/g. Table 2 summarizes the properties of the regenerated catalyst.Following regeneration, the catalyst was again tested in the mannerdescribed in Example 7. The catalysts before and after regeneration hadsimilar activities (i.e., DIPB conversion at a given temperature) andproduct purities (i.e., NPB/cumene at a given DIPB conversion) andtherefore indicate good catalyst regenerability.

EXAMPLE 9

A sample of the catalyst prepared in Example 4 was tested in the mannerdescribed in Example 7, as described previously. After testing, thespent catalyst was regenerated in the manner described in Example 8.Following regeneration, the catalyst was again tested in the mannerdescribed in Example 7.

FIGS. 3 and 4 graphically illustrate the test results for the catalystsbefore regeneration (labeled “Example 4”) and after regeneration(labeled “Example 9”). The results indicate that the catalysts beforeand after regeneration had similar activities (i.e., DIPB conversion ata given temperature) and product purities (i.e., NPB/cumene at a givenDIPB conversion) that were both better than the curves 101, 201 of FIGS.3, 4 respectively for the Example 1 catalyst, and therefore indicategood catalyst regenerability.

These examples show the benefits of high activity and product purity intransalkylating poly-alkylates to cumene attributed to catalystsprepared by the process disclosed herein.

1. An aromatic transalkylation catalyst comprising from about 5 wt % toabout 90 wt % of a modified LZ-210 zeolite on a volatile-free basis, andan alumina binder; the modified LZ-210 zeolite having a bulk Si/Al₂molar ratio ranging from about 6.5 to about 27; wherein 63% to 95% ofaluminum in the modified LZ-210 zeolite is framework aluminum.
 2. Thecatalyst of claim 1, wherein the modified LZ-210 zeolite has a bulkSi/Al₂ molar ratio ranging from about 6.5 to about
 23. 3. The catalystof claim 1 wherein the modified LZ-210 zeolite has an absolute intensityas measured by X-ray diffraction of at least
 50. 4. The catalyst ofclaim 1 wherein the modified LZ-210 zeolite has an absolute intensity asmeasured by X-ray diffraction of at least
 60. 5. The catalyst of claim 1wherein the modified LZ-210 zeolite has an absolute intensity asmeasured by X-ray diffraction of at least
 75. 6. The catalyst of claim1, wherein the modified LZ-210 zeolite has a Na₂O content of less than 3wt % based on the weight of the zeolite on a water-free basis.
 7. Thecatalyst of claim 1 wherein the catalyst has a loss on ignition (LOI) atabout 900° C. ranging from about 2 to about 4 wt %.
 8. The catalyst ofclaim 1, wherein the modified LZ-210 zeolite has a unit cell sizeranging from about 24.34 to about 24.58 Å.
 9. The catalyst of claim 1,wherein the catalyst comprises from about 50 wt % to about 85 wt % of amodified LZ-210 zeolite on a volatile-free basis.
 10. The catalyst ofclaim 1 wherein the catalyst is devoid of a metal hydrogenationcomponent.
 11. The catalyst of claim 1, further comprising a metalhydrogenation component in an amount less than 0.2 wt % of the catalyst.12. An aromatic transalkylation catalyst comprising from about 5 wt % toabout 90 wt % of a modified LZ-210 zeolite on a volatile-free basis, andan alumina binder; the modified LZ-210 zeolite having an absoluteintensity as measured by X-ray diffraction of at least 50; wherein 63%to 95% of aluminum in the modified LZ-210 zeolite is framework aluminum.13. The catalyst of claim 12, wherein a product of the absoluteintensity of the modified LZ-210 zeolite as measured by X-raydiffraction and the percentage of the aluminum of the modified LZ-210zeolite that is framework aluminum taken as a whole number is greaterthan
 4200. 14. The catalyst of claim 12 wherein the modified LZ-210zeolite has a bulk Si/Al₂ molar ratio ranging from about 6.5 to about 27and a unit cell size ranging from about 24.34 to about 24.58 Å.
 15. Aprocess for transalkylating aromatics, the process comprising passing atransalkylatable aromatic and an aromatic substrate to a reaction zone;contacting the transalkylatable aromatic and the aromatic substrate withthe catalyst of claim 1 in the reaction zone at transalkylationconditions including a temperature from about 100° C. to about 390° C.,a pressure of between 1 and 130 atmospheres, and a water concentrationof less than 20 wt-ppm based on the combined weight of thetransalkylatable aromatic and the aromatic substrate passed to thereaction zone.
 16. The process of claim 15 wherein transalkylationtemperature ranges from about 110° C. to about 275° C. and thetransalkylation pressure ranges from about 10 to about 40 atmospheres.17. The process of claim 15 wherein the transalkylatable aromaticcomprises at least one polyalkylaromatic selected from the groupconsisting of polyisopropylbenzenes and polyethylbenzenes.
 18. A processfor transalkylating aromatics, the process comprising passing atransalkylatable aromatic and an aromatic substrate to a reaction zone;contacting the transalkylatable aromatic and the aromatic substrate withthe catalyst of claim 12 in the reaction zone at transalkylationconditions including a temperature from about 100° C. to about 390° C.,a pressure of between 1 and 130 atmospheres, and a water concentrationof less than 20 wt-ppm based on the combined weight of thetransalkylatable aromatic and the aromatic substrate passed to thereaction zone.
 19. The process of claim 18 wherein transalkylationtemperature ranges from about 110° C. to about 275° C. and thetransalkylation pressure ranges from about 10 to about 40 atmospheres.20. The process of claim 18 wherein the transalkylatable aromaticcomprises at least one polyalkylaromatic selected from the groupconsisting of polyisopropylbenzenes and polyethylbenzenes.